This invention relates to a process for the separation of ethane and heavier hydrocarbons or propane and heavier hydrocarbons from liquefied natural gas (hereinafter referred to as LNG) combined with the separation of a gas containing hydrocarbons to provide a volatile methane-rich gas stream and a less volatile natural gas liquids (NGL) or liquefied petroleum gas (LPG) stream.
As an alternative to transportation in pipelines, natural gas at remote locations is sometimes liquefied and transported in special LNG tankers to appropriate LNG receiving and storage terminals. The LNG can then be re-vaporized and used as a gaseous fuel in the same fashion as natural gas. Although LNG usually has a major proportion of methane, i.e., methane comprises at least 50 mole percent of the LNG, it also contains relatively lesser amounts of heavier hydrocarbons such as ethane, propane, butanes, and the like, as well as nitrogen. It is often necessary to separate some or all of the heavier hydrocarbons from the methane in the LNG so that the gaseous fuel resulting from vaporizing the LNG conforms to pipeline specifications for heating value. In addition, it is often also desirable to separate the heavier hydrocarbons from the methane and ethane because these hydrocarbons have a higher value as liquid products (for use as petrochemical feedstocks, as an example) than their value as fuel.
Although there are many processes which may be used to separate ethane and/or propane and heavier hydrocarbons from LNG, these processes often must compromise between high recovery, low utility costs, and process simplicity (and hence low capital investment). U.S. Pat. Nos. 2,952,984; 3,837,172; 5,114,451; and 7,155,931 describe relevant LNG processes capable of ethane or propane recovery while producing the lean LNG as a vapor stream that is thereafter compressed to delivery pressure to enter a gas distribution network. However, lower utility costs may be possible if the lean LNG is instead produced as a liquid stream that can be pumped (rather than compressed) to the delivery pressure of the gas distribution network, with the lean LNG subsequently vaporized using a low level source of external heat or other means. U.S. Pat. Nos. 6,604,380; 6,907,752; 6,941,771; 7,069,743; and 7,216,507 and co-pending application Ser. Nos. 11/749,268 and 12/060,362 describe such processes.
Economics and logistics often dictate that LNG receiving terminals be located close to the natural gas transmission lines that will transport the re-vaporized LNG to consumers. In many cases, these areas also have plants for processing natural gas produced in the region to recover the heavier hydrocarbons contained in the natural gas. Available processes for separating these heavier hydrocarbons include those based upon cooling and refrigeration of gas, oil absorption, and refrigerated oil absorption. Additionally, cryogenic processes have become popular because of the availability of economical equipment that produces power while simultaneously expanding and extracting heat from the gas being processed. Depending upon the pressure of the gas source, the richness (ethane, ethylene, and heavier hydrocarbons content) of the gas, and the desired end products, each of these processes or a combination thereof may be employed.
The cryogenic expansion process is now generally preferred for natural gas liquids recovery because it provides maximum simplicity with ease of startup, operating flexibility, good efficiency, safety, and good reliability. U.S. Pat. Nos. 3,292,380; 4,061,481; 4,140,504; 4,157,904; 4,171,964; 4,185,978; 4,251,249; 4,278,457; 4,519,824; 4,617,039; 4,687,499; 4,689,063; 4,690,702; 4,854,955; 4,869,740; 4,889,545; 5,275,005; 5,555,748; 5,566,554; 5,568,737; 5,771,712; 5,799,507; 5,881,569; 5,890,378; 5,983,664; 6,182,469; 6,578,379; 6,712,880; 6,915,662; 7,191,617; 7,219,513; reissue U.S. Pat. No. 33,408; and co-pending application Ser. Nos. 11/430,412; 11/839,693; 11/971,491; and 12/206,230 describe relevant processes (although the description of the present invention is based on different processing conditions than those described in the cited U.S. patents).
The present invention is generally concerned with the integrated recovery of propylene, propane, and heavier hydrocarbons from such LNG and gas streams. It uses a novel process arrangement to integrate the heating of the LNG stream and the cooling of the gas stream to eliminate the need for a separate vaporizer and the need for external refrigeration, allowing high C3 component recovery while keeping the processing equipment simple and the capital investment low. Further, the present invention offers a reduction in the utilities (power and heat) required to process the LNG and gas streams, resulting in lower operating costs than other processes, and also offering significant reduction in capital investment.
Heretofore, assignee's co-pending application Ser. No. 12/060,362 could be used to recover C3 components and heavier hydrocarbon components in plants processing LNG, while assignee's U.S. Pat. No. 5,799,507 has been used to recover C3 components and heavier hydrocarbon components in plants processing natural gas. Surprisingly, applicants have found that by integrating certain features of the assignee's co-pending application Ser. No. 12/060,362 with certain features of the assignee's U.S. Pat. No. 5,799,507, extremely high C3 component recovery levels can be accomplished using less energy than that required by individual plants to process the LNG and natural gas separately.
A typical analysis of an LNG stream to be processed in accordance with this invention would be, in approximate mole percent, 92.2% methane, 6.0% ethane and other C2 components, 1.1% propane and other C3 components, and traces of butanes plus, with the balance made up of nitrogen. A typical analysis of a gas stream to be processed in accordance with this invention would be, in approximate mole percent, 80.1% methane, 9.5% ethane and other C2 components, 5.6% propane and other C3 components, 1.3% iso-butane, 1.1% normal butane, 0.8% pentanes plus, with the balance made up of nitrogen and carbon dioxide. Sulfur containing gases are also sometimes present.
FIGS. 1 and 2 are provided to quantify the advantages of the present invention.
In the following explanation of the above figures, tables are provided summarizing flow rates calculated for representative process conditions. In the tables appearing herein, the values for flow rates (in moles per hour) have been rounded to the nearest whole number for convenience. The total stream rates shown in the tables include all non-hydrocarbon components and hence are generally larger than the sum of the stream flow rates for the hydrocarbon components. Temperatures indicated are approximate values rounded to the nearest degree. It should also be noted that the process design calculations performed for the purpose of comparing the processes depicted in the figures are based on the assumption of no heat leak from (or to) the surroundings to (or from) the process. The quality of commercially available insulating materials makes this a very reasonable assumption and one that is typically made by those skilled in the art.
For convenience, process parameters are reported in both the traditional British units and in the units of the Système International d'Unités (SI). The molar flow rates given in the tables may be interpreted as either pound moles per hour or kilogram moles per hour. The energy consumptions reported as horsepower (HP) and/or thousand British Thermal Units per hour (MBTU/Hr) correspond to the stated molar flow rates in pound moles per hour. The energy consumptions reported as kilowatts (kW) correspond to the stated molar flow rates in kilogram moles per hour.
FIG. 1 is a flow diagram showing the design of a processing plant to recover C3+ components from natural gas using an LNG stream to provide refrigeration. In the simulation of the FIG. 1 process, inlet gas enters the plant at 126° F. [52° C.] and 600 psia [4,137 kPa(a)] as stream 31. If the inlet gas contains a concentration of sulfur compounds which would prevent the product streams from meeting specifications, the sulfur compounds are removed by appropriate pretreatment of the feed gas (not illustrated). In addition, the feed stream is usually dehydrated to prevent hydrate (ice) formation under cryogenic conditions. Solid desiccant has typically been used for this purpose.
The inlet gas stream 31 is cooled in heat exchanger 12 by heat exchange with a portion (stream 72a) of partially warmed LNG at −173° F. [−114° C.] and cool residue vapor stream 38. The cooled stream 31a enters separator 13 at −76° F. [−60° C.] and 584 psia [4,027 kPa(a)] where the vapor (stream 34) is separated from the condensed liquid (stream 35). Liquid stream 35 is flash expanded through an appropriate expansion device, such as expansion valve 17, to the operating pressure (approximately 450 psia [3,101 kPa(a)]) of fractionation tower 20. The expanded stream 35a leaving expansion valve 17 reaches a temperature of −88° F. [−67° C.] and is supplied to fractionation tower 20 at a first mid-column feed point.
The vapor from separator 13 (stream 34) enters a work expansion machine 10 in which mechanical energy is extracted from this portion of the high pressure feed. The machine 10 expands the vapor substantially isentropically to the tower operating pressure, with the work expansion cooling the expanded stream 34a to a temperature of approximately −96° F. [−71° C.]. The typical commercially available expanders are capable of recovering on the order of 80-88% of the work theoretically available in an ideal isentropic expansion. The work recovered is often used to drive a centrifugal compressor (such as item 11) that can be used to re-compress the heated residue vapor (stream 38a), for example. The expanded stream 34a is supplied to fractionation tower 20 at a second mid-column feed point.
The deethanizer in tower 20 is a conventional distillation column containing a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing to provide the necessary contact between the liquids falling downward and the vapors rising upward. The column also includes one or more reboilers (such as reboiler 19) which heat and vaporize a portion of the liquids flowing down the column to provide the stripping vapors which flow up the column to strip the liquid product, stream 41, of methane, C2 components, and lighter components. Liquid product stream 41 exits the bottom of the tower at 210° F. [99° C.], based on a typical specification of an ethane to propane ratio of 0.020:1 on a molar basis in the bottom product.
Overhead distillation stream 43 is withdrawn from the upper section of fractionation tower 20 at −87° F. [−66° C.] and is divided into two portions, streams 44 and 47. The first portion, stream 44, flows to reflux condenser 23 where it is cooled to −237° F. [−149° C.] and totally condensed by heat exchange with a portion (stream 72) of the cold LNG (stream 71a). Condensed stream 44a enters reflux separator 24 wherein the condensed liquid (stream 46) is separated from any uncondensed vapor (stream 45). The liquid stream 46 from reflux separator 24 is pumped by reflux pump 25 to a pressure slightly above the operating pressure of deethanizer 20 and stream 46a is then supplied as cold top column feed (reflux) to deethanizer 20. This cold liquid reflux absorbs and condenses the C3 components and heavier hydrocarbon components from the vapors rising in the upper section of deethanizer 20.
The second portion (stream 47) of overhead vapor stream 43 combines with any uncondensed vapor (stream 45) from reflux separator 24 to form cool residue vapor stream 38 at −88° F. [−67° C.]. Residue vapor stream 38 passes countercurrently to inlet gas in heat exchanger 12 where it is heated to −5° F. [−21° C.] (stream 38a). The residue vapor stream is then re-compressed in two stages. The first stage is compressor 11 driven by expansion machine 10. The second stage is compressor 21 driven by a supplemental power source which compresses stream 38b to sales line pressure (stream 38c). After cooling to 126° F. [52° C.] in discharge cooler 22, stream 38d combines with warm LNG stream 71b to form the residue gas product (stream 42). Residue gas stream 42 flows to the sales gas pipeline at 1262 psia [8,701 kPa(a)], sufficient to meet line requirements.
The LNG (stream 71) from LNG tank 50 enters pump 51 at −251° F. [−157° C.]. Pump 51 elevates the pressure of the LNG sufficiently so that it can flow through heat exchangers and thence to the sales gas pipeline. Stream 71a exits the pump 51 at −242° F. [−152° C.] and 1364 psia [9,404 kPa(a)] and is divided into two portions, streams 72 and 73. The first portion, stream 72, is heated as described previously to −173° F. [−114° C.] in reflux condenser 23 as it provides cooling to the portion (stream 44) of overhead vapor stream 43 from fractionation tower 20, and to 46° F. [8° C.] in heat exchanger 12 as it provides cooling to the inlet gas. The second portion, stream 73, is heated to 40° F. [4° C.] in heat exchanger 53 using low level utility heat. The heated streams 72b and 73a recombine to form warm LNG stream 71b, which thereafter combines with residue vapor stream 38d to form residue gas stream 42 as described previously.
A summary of stream flow rates and energy consumption for the process illustrated in FIG. 1 is set forth in the following table:
TABLE I(FIG. 1)Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr]StreamMethaneEthanePropaneButanes+Total3142,5455,0482,9721,65853,1453434,2891,7443134537,216358,2563,3042,6591,61315,9294349,0155,74720055,843446,470758307,3714500000466,470758307,3714742,5454,98917048,4723842,5454,98917048,4727140,2932,642491343,6897231,4292,061383234,077738,86458110819,6124282,8387,631508392,161410592,9551,6584,673Recoveries*Propane85.33%Butanes+99.83%PowerLNG Feed Pump3,561HP [5,854 kW]Reflux Pump21HP   [35 kW]Residue Gas Compressor21,779HP[35,804 kW]Totals25,361HP[41,693 kW]Low Level Utility HeatLNG Heater48,190MBTU/Hr[31,128 kW]High Level Utility HeatDemethanizer Reboiler108,000MBTU/Hr[69,762 kW]Specific PowerHP-Hr/Lb. Mole5.427[kW-Hr/kg mole][8.922]*(Based on un-rounded flow rates)
The recoveries reported in Table I are computed relative to the total quantities of propane and butanes+ contained in the gas stream being processed in the plant and in the LNG stream. Although the recoveries are quite high relative to the heavier hydrocarbons contained in the gas being processed (99.42% and 100.00%, respectively, for propane and butanes+), none of the heavier hydrocarbons contained in the LNG stream are captured in the FIG. 1 process. In fact, depending on the composition of LNG stream 71, the residue gas stream 42 produced by the FIG. 1 process may not meet all pipeline specifications. The specific power reported in Table I is the power consumed per unit of liquid product recovered, and is an indicator of the overall process efficiency.
FIG. 2 is a flow diagram showing processes to recover C3+ components from LNG and natural gas in accordance with co-pending application Ser. No. 12/060,362 and U.S. Pat. No. 5,799,507, respectively, with the processed LNG stream used to provide refrigeration for the natural gas plant. The processes of FIG. 2 have been applied to the same LNG stream and inlet gas stream compositions and conditions as described previously for FIG. 1.
In the simulation of the FIG. 2 process, the LNG to be processed (stream 71) from LNG tank 50 enters pump 51 at −251° F. [−157° C.] to elevate the pressure of the LNG to 1364 psia [9,404 kPa(a)]. The high pressure LNG (stream 71a) then flows through heat exchanger 52 where it is heated from −242° F. [−152° C.] to −50° F. [−45° C.] (stream 71b) by heat exchange with compressed vapor stream 83a from booster compressor 56 and distillation vapor stream 73. The heated and vaporized stream 71b enters work expansion machine 55 in which mechanical energy is extracted as the vapor is expanded substantially isentropically to a pressure of about 455 psia [3,135 kPa(a)] (the operating pressure of fractionation column 62). The work expansion cools the expanded stream 71c to a temperature of approximately −122° F. [−86° C.], before it is supplied to fractionation column 62 at an upper mid-column feed point.
Expanded stream 71c enters fractionation column 62 in the lower region of the absorbing section of fractionation column 62. The liquid portion of stream 71c commingles with the liquids falling downward from the absorbing section and the combined liquid proceeds downward into the stripping section of deethanizer 62 (which includes reboiler 61). The vapor portion of expanded stream 71c rises upward through the absorbing section and is contacted with cold liquid falling downward to condense and absorb the C3 components and heavier components.
A distillation liquid stream 72 is withdrawn from the lower region of the absorbing section in deethanizer 62 and is routed to heat exchanger 52. The distillation liquid stream is heated from −121° F. [−85° C.] to −50° F. [−45° C.], partially vaporizing stream 72a before it is returned as a lower mid-column feed to deethanizer 62, in the middle region of the stripping section.
A portion of the distillation vapor (stream 73) is withdrawn from the upper region of the stripping section of deethanizer 62 at −46° F. [−43° C.]. This stream is then cooled and partially condensed (stream 73a) in exchanger 52 by heat exchange with LNG stream 71a and distillation liquid stream 72 as described previously. The partially condensed stream 73a flows to reflux separator 64 at −104° F. [−76° C.]. The operating pressure of reflux separator 64 (452 psia [3,113 kPa(a)]) is slightly below the operating pressure of deethanizer 62 to provide the driving force which causes distillation vapor stream 73 to flow through heat exchanger 52 and into reflux separator 64, where the condensed liquid (stream 75) is separated from the uncondensed vapor (stream 74).
The liquid stream 75 from reflux separator 64 is pumped by pump 65 to a pressure slightly above the operating pressure of deethanizer 62, and the pumped stream 75a is then divided into two portions. One portion, stream 76, is supplied as top column feed (reflux) to deethanizer 62. This cold liquid reflux absorbs and condenses the C3 components and heavier components rising in the upper rectification region of the absorbing section of deethanizer 62. The other portion, stream 77, is supplied to deethanizer 62 at a mid-column feed position located in the upper region of the stripping section in substantially the same region where distillation vapor stream 73 is withdrawn, to provide partial rectification of stream 73. The deethanizer overhead vapor (stream 79) exits the top of deethanizer 62 at −105° F. [−76° C.] and is combined with the uncondensed vapor (stream 74) to form cold vapor stream 83 at −105° F. [−76° C.]. The liquid product stream 80 exits the bottom of the tower at 174° F. [79° C.], based on a typical specification of an ethane to propane ratio of 0.020:1 on a molar basis in the bottom product.
Cold vapor stream 83 flows to compressor 56 driven by expansion machine 55 to increase the pressure of stream 83a sufficiently so that it can be totally condensed in heat exchanger 52. Stream 83a exits the compressor at −58° F. [−50° C.] and 669 psia [4,611 kPa(a)] and is cooled to −114° F. [−81° C.] (stream 83b) by heat exchange with the high pressure LNG feed stream 71a and distillation liquid stream 72 as discussed previously. Condensed stream 83b is pumped by pump 63 to a pressure slightly above the sales gas delivery pressure for subsequent vaporization in heat exchangers 23 and 12, heating stream 83c from −94° F. [−70° C.] to 40° F. [4° C.] as described in paragraphs [0033] and [0037] below to produce warm lean LNG stream 83e. 
In the simulation of the FIG. 2 process, inlet gas enters the plant at 126° F. [52° C.] and 600 psia [4,137 kPa(a)] as stream 31. The feed stream 31 is cooled in heat exchanger 12 by heat exchange with cool lean LNG (stream 83d) at −56° F. [−49° C.], cool residue vapor stream 38, and separator liquids (stream 35a). The cooled stream 31a enters separator 13 at −51° F. [−46° C.] and 584 psia [4,027 kPa(a)] where the vapor (stream 34) is separated from the condensed liquid (stream 35).
The vapor from separator 13 (stream 34) enters a work expansion machine 10 in which mechanical energy is extracted from this portion of the high pressure feed. The machine 10 expands the vapor substantially isentropically to the operating pressure of fractionation tower 20 (approximately 441 psia [3,039 kPa(a)]), with the work expansion cooling the expanded stream 34a to a temperature of approximately −73° F. [−58° C.]. The partially condensed expanded stream 34a is then supplied as feed to fractionation tower 20 at an upper mid-column feed point. The liquid portion of stream 34a commingles with the liquids falling downward from the absorbing section and the combined liquid proceeds downward into the stripping section of deethanizer 20 (which includes reboiler 19). The vapor portion of expanded stream 34a rises upward through the absorbing section and is contacted with cold liquid falling downward to condense and absorb the C3 components and heavier components.
Liquid stream 35 is flash expanded through an appropriate expansion device, such as expansion valve 17, to slightly above the operating pressure of fractionation tower 20. The expanded stream 35a leaving expansion valve 17 reaches a temperature of −62° F. [−52° C.] before it provides cooling to the incoming feed gas in heat exchanger 12 as described previously. The heated stream 35b at 82° F. [28° C.] then enters fractionation tower 20 at a lower mid-column feed point to be stripped of its methane and C2 components.
A distillation liquid stream 36 is withdrawn from the lower region of the absorbing section in deethanizer 20 and is routed to heat exchanger 23. The distillation liquid stream is heated from −86° F. [−66° C.] to −12° F. [−24° C.], partially vaporizing stream 36a before it is returned as a lower mid-column feed to deethanizer 20, in the middle region of the stripping section.
A portion of the distillation vapor (stream 37) is withdrawn from the upper region of the stripping section of deethanizer 20 at −9° F. [−23° C.]. This stream is then cooled and partially condensed (stream 37a) in exchanger 23 by heat exchange with cold lean LNG stream 83c and with distillation liquid stream 36 as described previously. The partially condensed stream 37a flows to reflux separator 24 at −86° F. [−65° C.]. The operating pressure of reflux separator 24 (437 psia [3,012 kPa(a)]) is slightly below the operating pressure of deethanizer 20 to provide the driving force which causes distillation vapor stream 37 to flow through heat exchanger 23 and into reflux separator 24, where the condensed liquid (stream 45) is separated from the uncondensed vapor (stream 44).
The liquid stream 45 from reflux separator 24 is pumped by pump 25 to a pressure slightly above the operating pressure of deethanizer 20, and the pumped stream 45a is then divided into two portions. One portion, stream 46, is supplied as top column feed (reflux) to deethanizer 20. This cold liquid reflux absorbs and condenses the C3 components and heavier components rising in the upper rectification region of the absorbing section of deethanizer 20. The other portion, stream 47, is supplied to deethanizer 20 at a mid-column feed position located in the upper region of the stripping section in substantially the same region where distillation vapor stream 37 is withdrawn, to provide partial rectification of stream 37.
The deethanizer overhead vapor (stream 43) exits the top of deethanizer 20 at −88° F. [−67° C.] and is directed into heat exchanger 23 to provide cooling to distillation vapor stream 36 as described previously. The heated overhead vapor stream 43a at −56° F. [−49° C.] is combined with the uncondensed vapor (stream 44) to form cool residue vapor stream 38 at −58° F. [−50° C.]. The liquid product stream 40 exits the bottom of the tower at 208° F. [98° C.], based on a typical specification of an ethane to propane ratio of 0.020:1 on a molar basis in the bottom product.
Cool residue vapor stream 38 passes countercurrently to inlet gas stream 31 in heat exchanger 12 where it is heated to 8° F. [−13° C.] (stream 38a). The heated residue vapor stream is then re-compressed in two stages. The first stage is compressor 11 driven by expansion machine 10. The second stage is compressor 21 driven by a supplemental power source which compresses stream 38b to sales line pressure (stream 38c). After cooling to 126° F. [52° C.] in discharge cooler 22, stream 38d combines with warm lean LNG stream 83e to form the residue gas product (stream 42). Residue gas stream 42 flows to the sales gas pipeline at 1262 psia [8,701 kPa(a)], sufficient to meet line requirements.
A summary of stream flow rates and energy consumption for the process illustrated in FIG. 2 is set forth in the following table:
TABLE II(FIG. 2)Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr]StreamMethaneEthanePropaneButanes+Total3142,5455,0482,9721,65853,1453438,3512,82068611442,843354,1942,2282,2861,54410,302364,6514,42079211410,0373712,89411,068217124,339443,255403203,705459,63910,665215120,634465,5916,186125111,968474,0484,4799008,6664339,2904,58619044,7713842,5454,98921048,476400592,9511,6584,6697140,2932,642491343,6897211,7402,966264115,0007331,07910,63159041,8357414,9839911016,0237516,0969,64058025,812768,0484,82029012,906778,0484,82029012,9067925,3101,6413027,1668340,2932,6324043,1898001048735004282,8387,62125091,665410693,4381,6615,169Recoveries*Propane99.29%Butanes+100.00%PowerLNG Feed Pump3,552HP [5,839 kW]LNG Product Pump2,766HP [4,547 kW]Reflux Pump 2580HP  [132 kW]Reflux Pump 6396HP  [158 kW]Residue Gas Compressor22,801HP[37,485 kW]Totals29,295HP[48,161 kW]High Level Utility HeatDeethanizer Reboiler 1957,670MBTU/Hr[37,252 kW]Deethanizer Reboiler 6199,590MBTU/Hr[64,330 kW]Totals157,260MBTU/Hr[101,582 kW] Specific PowerHP-Hr/Lb. Mole5.667[kW-Hr/kg mole][9.317]*(Based on un-rounded flow rates)
Comparison of the recovery levels displayed in Tables I and II shows that the liquids recovery of the FIG. 2 processes is higher than that of the FIG. 1 process due to the recovery of the heavier hydrocarbon liquids contained in the LNG stream in fractionation tower 62. The propane recovery improves from 85.33% to 99.29% and the butanes+ recovery improves from 99.83% to 100.00%. The process efficiency of the FIG. 2 processes is slightly lower, however, about 4% in terms of the specific power relative to the FIG. 1 process.